Synthesis gas (or “syngas”) is a mixture of primarily hydrogen and carbon monoxide, commonly with carbon dioxide, methane, water, nitrogen, and possibly other constituents such as argon or helium. Syngas can be produced from any carbon containing feedstock, including natural gas, heavy petroleum cuts such as resid or coke, bitumen, coal, or biomass by a variety of processes. Natural gas may be converted to synthesis gas using steam methane reforming (“SMR”), carbon dioxide reforming (“dry reforming”), or a combination of these two processes, or by processes employing air, enriched air, or oxygen (generally with some steam addition) such as combined reforming, autothermal reforming (“ATR”), catalytic partial oxidation (“CPOX”) or thermal partial oxidation (“POX”). When employed, enriched air or oxygen may be produced by conventional cryogenic air separation, pressure swing adsorption (“PSA”) or membrane processes, the use of ion transport membranes (“ITM”) or any other method producing a gas containing sufficient oxygen. Heavier liquids or solids, such as petroleum cuts and/or coal, are typically converted to synthesis gas by processes employing air, enriched air, or oxygen such as gasification and catalytic/thermal partial oxidation, again typically with some steam addition to the reaction mixture. Numerous synthesis gas production processes are well known in the art.
As used herein, the term “catalyst” covers both the catalyst in active form and the catalyst in precursor form. Iron, cobalt, nickel, and ruthenium are active Fischer-Tropsch catalyst metals. Nickel-containing catalysts typically produce a light product slate comprised primarily of C1-C4 gases. Therefore, nickel-containing catalysts are not generally practical for Fischer-Tropsch processes when the desired product slate is comprised primarily of C5+ liquids. Ruthenium-containing catalysts generally exhibit very good to excellent selectivity for heavier (i.e. C5+) hydrocarbon products, but may be too rare and expensive for practical use. Iron is typically employed as an FT catalyst in precipitated or sintered (i.e., unsupported) form, with a variety of chemical (e.g., potassium, copper) and/or structural (e.g., alumina, silica) promoters added. Cobalt, for use in an FT catalyst, is typically dispersed on a porous refractory oxide support such as alumina, silica, titania, zinc oxide, or a zeolite, with the addition of reduction promoters (typically a platinum group metal (“PGM”) such as ruthenium, rhenium, platinum, or palladium) and other chemical and/or support modifiers (e.g., rare earth oxides, alkalis). Iron- and cobalt-containing FT catalyst compositions and formulations are well known in the art. The product of a Fischer-Tropsch synthesis reaction, irrespective of catalyst used, is referred to herein as “synthetic crude.”
Fischer-Tropsch Synthesis, Fischer-Tropsch Catalyst Activation and Regeneration
Both iron- and cobalt-containing catalysts must be activated prior to FT synthesis operations. As used herein, the term “activation” means activation, reductions or activation/reduction. Iron-based catalysts may be activated using hydrogen, hydrogen-carbon monoxide containing mixtures (i.e., synthesis gases), or carbon monoxide. Activation conditions are fairly mild, with temperatures in the range of 200-300° C., although temperatures up to about 400° C. have been employed in the past. The extent of reduction for iron-based catalysts need not be particularly high to achieve effective activation of the iron-based catalysts. In contrast, cobalt-containing catalysts are reduced with hydrogen gas, mixtures of hydrogen gas with inert gases such as helium or argon, or mixtures of hydrogen with small amounts of hydrocarbon gases, at temperatures in the range of 300-400° C., and the extent of reduction must be relatively high, usually more than 70-80%. For both iron and cobalt the concentration of water (which is a product of the metal reduction) in the activation vessel is generally limited or controlled, which typically involves both cooling/condensation and the use of water absorbents or dryers. In general, nickel- and ruthenium-based catalysts are more similar to cobalt-based catalysts in behavior and conditions for activation than they are to iron-based catalysts, although ruthenium-based catalysts may not require reduction promoters.
Procedures for activation of non-iron based FT catalysts require the addition of systems and equipment over that necessary for syngas conversion operations. Such systems and equipment include, for example, that needed for generation of relatively pure hydrogen, recycle compressors to minimize H2 (and therefore generation rates), as well as high temperature heaters. Procedures for activation of iron based FT catalysts may also require system and equipment for relatively high concentration carbon monoxide generation and recycle. In addition, the design temperatures for the equipment in which the activation processes occur are, in general, significantly higher than that necessary for the syngas conversion processes.
For FT synthesis, iron-containing catalysts are normally employed at temperatures in the 180-300° C. range as 1-3 mm particulate particles in multi-tubular fixed bed reactors, or as 10-500 μm particles in three-phase slurry bubble column reactors (low temperature FT processes) or at temperatures in the 300-350° C. range as class “A” powders in circulating or fixed fluidized bed reactors (high temperature FT processes). Cobalt (and nickel or ruthenium) catalysts can be employed in multi-tubular fixed bed reactors, three-phase slurry bubble column reactors, and (typically as coatings on) monoliths, tube wall reactors, or microchannel reactors, all at relatively low temperatures, 180-250° C. FT-based plants employ either single stage FT reactor(s) using tail gas recycle or multiple stages once-through to achieve high overall CO conversion while limiting per pass conversions to the 50-70% range.
Additional equipment and design parameters and conditions required for FT catalyst activation processes are known in the art and are readily apparent from U.S. Pat. Nos. 4,525,206; 4,585,798; 4,605,676; 4,605,679; 4,670,414; 4,670,475; 4,717,702; 4,729,981; 4,801,573; 4,857,497; 4,962,078; 5,585,316; 6,706,661; 6,716,886; 6,753,351 and U.S. Application Publication Nos. 20040242941; 20050250863; 20070004810; 20070112080, all of which are incorporated by reference in their entirety as part of the disclosure herein.
Following activation, FT catalysts are generally pyrophoric. Thus, if the FT catalyst is not activated within the FT reactor, or at least activated locally and loaded into the reactor shortly thereafter, it is generally passivated to ease intermediate handling. Techniques for passivating the various metal-based FT catalysts as well as the equipment needed for such passivation are known in the art. U.S. Pat. Nos. 4,607,020; 4,717,702; 4,729,981; 4,863,894; 5,756,419; 5,928,985; 6,777,451; and 6,815,388 all of which are incorporated herein by reference in their entirety as part of the disclosure herein, disclose techniques and associated equipment, parameters and conditions for passivating Co, Ni, Cu, Ru, Rh, Pd, Os, Ir, Re, Pt and Fe based catalysts, including FT and/or oxygenate synthesis catalysts.
In the simplest form the FT reaction may be written as:(n)CO+(2n)H2→—(CH2)n-+(n)H2Owhere “n” ranges from 1 to as high as several hundred. The relationship between the amounts of the different carbon number (“N”) products can usually be described in terms of a polymerization probability alpha (“α”) with values between 0 and 1, which is written as:α=Cn+1/Cn, where Cn and Cn+1 are in mole percents.Thus a plot of the natural logarithm of the mole percent of each carbon number product versus the carbon number will yield a straight line with slope −α. The products of low temperature FT processes are predominantly (>>50 mole %) normal paraffins, although olefins, alcohols, and branched paraffins may be present in smaller amounts, up to 5-10 mole %.
In addition to the FT reaction described above the water-gas-shift (“WGS”) reaction may also occur. This can be written as:CO+H2O→H2+CO2 
Iron typically has a higher activity for the WGS reaction than other FT catalytic metals; depending on the exact feed gas composition and reaction conditions, the selectivity to CO2 on iron catalysts will typically range from 8-12% to as high as 40%. Except under very unusual or high (>80-85%) CO conversion conditions the CO2 selectivity with cobalt catalysts is typically in the 0-2% range. The high WGS activity of iron is potentially advantageous when converting synthesis gases with H2/CO ratios lower than the stoichiometric consumption ratio of about 2.05-2.10, such as those produced from coal, coke, or biomass gasification/partial oxidation (where the H2/CO ratio of the produced synthesis gas can be as low as 0.6), as additional H2 is generated in the FT reactor. The use of cobalt catalysts with such H2 deficient synthesis gas feeds requires the addition of a separate (external and upstream) WGS reactor and catalyst to increase the H2/CO ratio to the stoichiometric consumption ratio; otherwise CO conversion is debited by the lack of H2 and severe catalyst deactivation due to surface carbon formation may occur.
For synthesis gas feeds with H2/CO ratios in the range of 2.0 and higher, such as those produced from natural gas steam-methane or autothermal reforming or partial oxidation, the relatively high WGS activity of iron catalysts is a significant debit, as additional CO is converted to H2 (and CO2), which is generally not needed. Non-shifting (i.e., cobalt, ruthenium, or nickel) catalysts are strongly favored for converting natural gas based synthesis gases.
Regardless of the catalyst type and reaction conditions, as well as the reactor type, FT catalysts deactivate with time on stream. FT catalyst deactivation mechanisms can be broadly divided into two categories: those due to external causes, and those due to internal causes. External causes include the presence of catalyst poisons (contaminants) in the feed synthesis gas. These potential poisons include sulfur, mercury, arsenic, halides (especially chlorine), ammonia, and hydrogen cyanide. Internal causes include: (1) the build-up of high molecular weight (“MW”) wax in the catalyst pores; (2) self-poisoning by-product organic acids; (3) build-up of hydrogen deficient, coke pre-cursor carbon deposits and/or Boudouard carbon on the catalytic metals (especially at high(er) temperature); (4) loss of catalytic metal surface area due to hydrothermal sintering; (5) oxidation of the active metals to less active or inactive oxides; (6) formation of unreactive mixed metal compounds due to reaction between the active metal and inactive support (e.g., cobalt aluminates, silicates, and/or titanates); and (7) degradation of the support due to chemical, thermal, or mechanical means such as pore collapse, material leaching, or loss of structural integrity (e.g., crushing). The rate of catalyst deactivation can vary enormously, with catalyst life ranging from several days to several years depending on reaction conditions and feed poison concentrations. The various mechanisms of FT catalyst deactivation are known in the art.
As catalyst deactivation proceeds at some rate the activity will typically reach a lower limit at which it is no longer practical to continue operations. This may occur when CO conversion (at constant feed rate and temperature), or feed rate (at constant conversion and temperature)—in other words total plant production—has dropped too low to be economic, or because temperature (at constant feed rate and conversion) has been increased to the maximum consistent with the desired product slate (i.e. maximum light hydrocarbon gas selectivity) and absence of rapid carbon formation on the catalyst surface. At such a time the catalyst can either be removed from the reactor and replaced with fresh catalyst or regenerated. Removal and replacement is more common for iron-based catalysts, which are relatively cheap but difficult to regenerate. Nevertheless, in some instances iron-based catalysts are regenerated. For example, U.S. Pat. No. 7,303,731, the disclosure of which is incorporated herein in its entirety, discloses a method for regenerating iron-based Fischer-Tropsch catalysts. Catalyst regeneration is more common for cobalt-based, as well as the less frequently used ruthenium- and nickel-based catalysts, which are generally more expensive than iron-based catalysts and easier to regenerate. After catalyst replacement (e.g., replacement of an iron-based FT catalyst), the appropriate catalyst activation procedure is employed. In contrast, where the FT catalyst is regenerated, regeneration procedures vary from activation-type processes to more complicated, multi-step procedures. A number of FT catalyst regeneration processes are known in the art, including those discussed in U.S. Pat. Nos. 4,207,208; 4,822,824; 5,128,377; 5,260,239; 5,268,344; 5,283,216; 5,288,673; 5,389,690; 5,817,701; 5,844,005; 2,487,867; 6,022,755; 6,162,754; 6,201,030; 6,486,220; 6,878,655; 6,949,488; 4,595,703, 1986; 4,663,305; 4,670,475; 4,738,948; 4,755,536; 4,822,824; 6,300,268; 6,455,596; 6,537,945; 6,753,286; 6,753,354; 6,800,579; 6,812,179; 6,838,487; 6,869,978; 6,900,151; 6,962,947; 6,989,403; and 2,289,731, which are all incorporated in their entirety into the disclosure herein.
Like activation processes, FT catalyst regeneration processes require additional equipment not otherwise needed for normal FT synthesis operations. In the simplest regeneration procedure—hydrogen treatment only—such additional equipment would be similar to that required for catalyst activation. Additional gas sources, vessels, and/or solids handling equipment can be required as the regeneration process complexity increases, especially for slurry bubble column reactor based systems where the small catalyst particles must be separated from the reactor liquid phase (i.e., typically the heavier fractions of the product wax). Both activation and regeneration processes may be conducted in discrete batches or continuously. Where activation and/or regeneration are conducted in discrete batches, some non-trivial level of manual operator intervention, action, and/or control is needed.
Most FT plants—existing or proposed, based on natural gas (“GTL”) or coal (“CTL”) feedstock—are in the size range between about 10,000 to more than 100,000 bbl/day. As used herein bbl/day capacities refer to either the FT reactor effluent or upgraded product, as upgrading process losses tend to be small, in the range of 5%. Such plants are currently land-based and in a fixed location. More recently proposed FT plants based on biomass feedstocks (“BTL”) are generally smaller, in the several thousand bbl/day range based on local feedstock availability and transportation costs, and are also currently exclusively land-based.
There are also a number of generally smaller (1,000 to perhaps 15-20,000 bbl/day) proposed offshore GTL plants located on ships, FPSO's, barges, and/or drilling rigs/production platforms. There have also been a number of proposed very small (“micro”) land-based FT plants, typically having 30-500 bbl/day production capacity, most of which are designed to be transportable, for example, by truck, rail, and/or helicopter. Plants in the higher end of this production range, i.e., >250 bbl/day, would require multiple modules, and are transportable with specialized means, while those at the lowest end, i.e., <<250 bbl/day, would likely fit in a single standard shipping container and could be relocated fairly easily, while the number of containers necessary for units in the middle of this range would depend on the specifics of the process design and hardware.
Oxygenate Synthesis, Oxygenate Synthesis Catalyst Activation and Regeneration
As used herein, the term oxygenate means any of (1) ethers, including, for example, ethyl tert-butyl ether (ETBE), diisopropyl ether (DIPE), dimethyl ether (DME), methyl tert-butyl ether (MTBE), tert-amyl ethyl ether (TAEE), tert-amyl methyl ether (TAME); and (2) methanol and C2+ alcohols, including for example, ethanol (EtOH), propanol, butanol, tert-amyl alcohol (TAA), and tert-butyl alcohol (TBA). Methanol is the most common oxygenated product synthesized from syngas, while the most common catalysts for methanol synthesis are based on copper. Catalysts for higher (C2+) oxygenate synthesis from synthesis gas are typically alkali promoted metals such as copper, zinc, molybdenum, chromium, palladium, cobalt, or rhodium, as well as mixtures of these various metals, especially cobalt and copper. Copper-based oxygenate synthesis catalysts are activated/reduced with hydrogen, much like cobalt-based FT catalysts.
Although there are a number of different oxygenate synthesis processes currently known in the art (especially for methanol synthesis), the majority of such processes utilize fixed bed reactors, whether tray, tube-shell, annular, radial, or microchannel. However, three-phase slurry methanol and/or dimethyl ether (DME) synthesis are also known.
The reactions involved in methanol synthesis can generally be described as:CO+2H2→CH3OHCO2+3H2→CH3OH+H2O;andCO+H2O→CO2+H2 (WGS reaction)A stoichiometric ratio, defined as (H2—CO2)/(CO+CO2), of somewhat higher than 2 is preferred. In methanol synthesis processes, by-products include water and fusel oil (may include acetone, formaldehyde, ethanol, methyl formate, mimethyl ether, methylal, methyl acetate, acetaldehyde, amyl and other higher alcohols).
For higher alcohol synthesis the general reaction can be described as:nCO+2nH2→CnH2n+1OH+(n−1)H2Owhere n is typically between 1 and about 8. Higher alcohol synthesis is generally less selective than methanol production, and a large number of competing reactions also occur. These include water-gas-shift, Fischer-Tropsch synthesis, and methanol synthesis.
Copper-based oxygenate synthesis catalysts must be activated with hydrogen, much like cobalt-based FT catalysts. In the case of copper-based oxygenate synthesis catalysts, however, the maximum reduction temperature is about 300° C. (being similar to or only somewhat higher than the usual maximum operating temperature range of about 250 to perhaps 300-350° C.). Non-copper-based oxygenate synthesis catalysts generally require activation at a minimum of 300° C., and many require temperatures as high as 350-400° C., similar to the temperature range required for activation of cobalt-based FT catalysts.
Like FT synthesis catalysts, oxygenate synthesis catalysts deactivate with time. Poisoning by feed contaminants is the primary mechanism for rapid deactivation; potential feed poisons being the same or similar to FT synthesis catalyst poisons, as is well known in the art. Sulfided oxygenate synthesis catalysts, developed primarily for higher alcohol synthesis, are typically somewhat resistant to sulfur poisoning. Methanol synthesis catalyst deactivation also occurs from thermal sintering (i.e., loss of catalytic surface area), especially at higher temperatures, and support degradation.
Methanol synthesis catalysts are typically not regenerated. If feed poison concentrations are minimized and excessive catalyst/bed temperatures are avoided conventional copper-based methanol synthesis catalysts can have a two to five year process life. However methanol synthesis catalyst regeneration procedures which are nearly identical to the hydrogen and/or oxygen-hydrogen regeneration procedures employed for FT catalyst regeneration are known.
Most existing oxygenate synthesis plants, mainly producing methanol, are smaller than FT synthesis plants, generally producing in the range of 500 to about 10,000 bbl/day FT liquids equivalent (300 million gallons per year methanol). Larger capacity oxygenate synthesis plants are also known in the art, with capacities of greater than 20,000 bbl/day. While most methanol synthesis plants use natural gas derived syngas, it is also known in the art to use coal, biomass, petroleum and/or other carbon/hydrocarbon source derived syngas. In addition, proposals for the conversion of off-shore stranded natural gas to methanol, using ship-, barge-, and/or platform-based methanol synthesis units are known.
Conventional Plant Configuration of Catalyst Activation and Regeneration
For a fixed land-based synthesis gas conversion plant (i.e., >500 bbl/day), there are a limited number of options for the location and method of the FT or oxygenate synthesis catalyst activation and regeneration facilities and processes. At one end of the spectrum is the initial shipment of new, unactivated (i.e., unreduced) catalyst from a catalyst manufacturing facility to an FT or oxygenate synthesis production plant, with the completely spent catalyst shipped out of the production plant facility (for metals reclamation and/or disposal) at the end of the catalyst's useful life. In this configuration, all required activation and/or regeneration facilities (including, for example, hydrogen generation and recovery, process gas and reactor heaters, recycle compressors, gas dryers and dryer regeneration equipment, powdered catalyst+wax mixing/slurrying devices, catalyst-wax separation devices, nitrogen and/or dilute air generation/recovery facilities) are located at the FT or oxygenate synthesis plant. At the other end of the spectrum, activated catalyst is shipped to the FT or oxygenate synthesis plant with no activation or regeneration processes conducted at the FT or oxygenate synthesis plant. In such a situation, spent catalyst may be removed from the FT or oxygenate synthesis plant and shipped to another location for regeneration, metals reclamation, and/or disposal.
There are, of course, a number of conceivable intermediate configurations between these extremes. For example, the initial charge of a slurry FT catalyst may be activated as part of the manufacturing procedure, then coated in FT wax and allowed to cool/solidify. The pre-reduced, wax encased catalyst may then be shipped to the FT or oxygenate synthesis plant and loaded into the FT or oxygenate synthesis reactor. Regardless of whether activation and/or regeneration facilities exist at the FT or oxygenate synthesis plant, pre-reduction of the initial charge of catalyst may be preferable in order to avoid commissioning local facilities long before the completion of the rest of the plant is scheduled.
Slurry FT catalyst hydrogen treatments require far less “extra” processing equipment and utilities than the more complicated oxidation/re-reduction regeneration processes. In some instances, it may be advantageous to conduct such hydrogen treatments at the FT or oxygenate synthesis plant, while shipping catalyst to another location for oxidation/reduction type regeneration. Currently, for FT or oxygenate synthesis plants, the catalyst manufacturing, and activation and/or regeneration facilities are in fixed locations, while catalyst is transported between them. As shown in FIG. 1 illustrating one currently known process, each of the catalyst manufacturing facility, catalyst activation and/or regeneration facility, and synthesis plant are remotely located from each other. Raw catalyst is transported from the Catalyst Manufacturing facility to the Activation and/or regeneration facility. Activated and/or regenerated catalyst is transported from the Activation and/or Regeneration Facility to the Synthesis Plant and spent catalyst is transported in the opposite direction. Spent catalyst may also be transported from the Synthesis Plant and/or Catalyst Regeneration Facility to the Catalyst Manufacturing facility for reclamation of metals (shown in FIGS. 1-3 as dashed lines). In an alternate embodiment, the Synthesis Plant is located remotely from the Catalyst Manufacturing and Activation and/or Regeneration locations. Referring to FIG. 2, a current process is shown in which the Synthesis Plant and Catalyst Activation and/or Regeneration Facility are located locally to each other but remotely from the Catalyst Manufacturing facility. In such situation, raw catalyst is transported from the Catalyst Manufacturing facility to the Catalyst Activation facility. Activated and/or regenerated catalyst is transported to the Synthesis Plant from the Catalyst Activation and/or Regeneration facility and spent catalyst is transported in the opposite direction. Spent catalyst may also be transported from either the Synthesis Plant or the Catalyst Regeneration facility to the Catalyst Manufacturing facility for metals reclamation. FIG. 3 illustrates a currently known process in which the Synthesis Plant and Catalyst Regeneration facility are located locally to each other but remotely from the Catalyst Manufacturing and Activation Facilities. As shown in FIG. 3, the catalyst, spent, activated and regenerated is transported between stationary facilities and plants. The distances between these facilities varies from a few miles (or less—if located at the same facility) to thousands of miles, depending primarily on whether the activation and/or regeneration facilities are located at the FT and/or oxygenate synthesis production facility (“synthesis plant”), and whether the synthesis plant is located in a generally developed, commercial area. As seen in FIGS. 1-3, in all known situations, however, the synthesis plant and the catalyst manufacturing, activation and/or regeneration facilities are in fixed locations, while catalyst is transported between them.
The more difficult and/or time consuming the removal of synthesis catalyst from the synthesis reactors is, the more likely the preferred configuration is to perform activation or regeneration treatments in the synthesis reactor itself. Catalyst addition and removal is easiest in the fluidized and slurry bed reactors, difficult but not impossible in particulate catalyst fixed-bed reactors, and extremely difficult in monolith, tube wall, or microchannel reactors with catalyst coatings. Moreover, the more frequently the synthesis catalyst must be activated and/or regenerated, the more desirable it is to conduct such activation and/or regeneration locally at the FT and/or oxygenate synthesis production plant.
Movable Plant Configuration of Catalyst Activation and Regeneration
For currently-proposed movable, predominantly offshore, plants practical options for the location and method of the FT or oxygenate synthesis catalyst activation and regeneration facilities are much the same as for conventional land-based synthesis plants. A preferred solution, however, is shifted strongly towards locating complicated, infrequently used equipment—such as that required for the oxidation/reduction-type catalyst regenerations—to a location other than the synthesis production plant, such as the catalyst manufacturing plant. Such preference is due to the significant plot plan limitations and/or costs associated with increased plot plan in movable facilities, e.g., FPSOs and barges. In addition, the production capacity of proposed movable synthesis plants are significantly less than that of conventional, land-based plants. Therefore, equipment associated with catalyst activation and/or regeneration constitutes a larger fraction of the total plant capital cost.